Method for converting hydrocarbon feedstock comprising a shale oil by decontamination, hydroconversion in an ebullating bed, and fractionation by atmospheric distillation

ABSTRACT

Method and plant for converting hydrocarbon feedstock comprising a shale oil, comprising a step of decontaminating, a step of hydroconverting in an ebullating bed, a step of fractionating into a light fraction, a naphtha fraction, a gas-oil fraction and a fraction heavier than the gas-oil fraction, the naphtha and gas-oil fractions being hydrotreated, the fraction heavier than the gas-oil fraction being conveyed to the decontaminating step. The method aims to maximize the yield of fuel bases.

The invention relates to a method for converting hydrocarbon feedstocks comprising a shale oil into lighter products which can be utilized as fuels and/or raw materials for petrochemistry. The invention relates more particularly to a method for converting hydrocarbon feedstocks comprising a shale oil that comprises a decontaminating step, a step of hydroconverting the decontaminated oil in an ebullating bed, followed by a step of fractionating by atmospheric distillation to give a light fraction, naphtha fraction and gas-oil fraction and to give a fraction heavier than the gas-oil fraction, and a dedicated hydrotreating for each of the naphtha and gas-oil fractions. This method enables shale oils to be converted into very-high-quality fuel bases, and is aimed more particularly at an excellent yield.

In view of high barrel price volatility and a reduction in discoveries of conventional petroleum fields, petroleum groups are turning towards non-conventional sources. Next to petroleum-bearing sands and deep offshore, bituminous shales, although relatively poorly known, are becoming ever more coveted.

Bituminous shales are sedimentary rocks which contain an insoluble organic substance called kerogen. By heat treatment in situ or ex situ (“retorting”) in the absence of air at temperatures of between 400 and 500° C., these shales liberate an oil, shale oil, with a general appearance like that of crude petroleum.

Although of a different composition from crude petroleum, shale oils may constitute a substitute for the latter and also a source of chemical intermediates.

Shale oils cannot be directly substituted into the applications of crude petroleum. Indeed, although these oils resemble petroleum in certain respects (for example, in a similar H/C ratio), they differ in their chemical nature and in a much greater level of metallic and/or non-metallic impurities, thereby making the converting of this non-conventional resource much more complex than that of petroleum. Shale oils have, in particular, levels of oxygen and of nitrogen that are much higher than those in petroleum. They may also contain higher concentrations of olefins, of sulphur or of metal compounds (especially compounds containing arsenic).

Shale oils obtained by pyrolysis of kerogen contain a large number of olefinic compounds resulting from cracking, and this translates into additional hydrogen demand at the refining stage. For instance, the bromine index, which enables calculation of the concentration by weight of olefinic hydrocarbons (by addition of bromine to the ethylenic double bond), is generally greater than 30 g/100 g of feedstock for shale oils, whereas it is between 1 and 5 g/100 g of feedstock for residues of petroleum. The olefinic compounds resulting from cracking are essentially composed of monoolefins and diolefins. The unsaturations present in the olefins are a potential source of instability by polymerization and/or oxidation.

The oxygen content is generally higher than in heavy crudes, and may be as much as 8% by weight of the feedstock. The oxygen compounds are often phenols or carboxylic acids. Consequently, shale oils may have a marked acidity.

The sulphur content varies between 0.1% and 6.5% by weight, necessitating stringent desulphurizing treatments in order to meet the specifications for fuel bases. The sulphur compounds are in the form of thiophenes, sulphides or disulphides. Moreover, the sulphur distribution profile within a shale oil may be different from that obtained in a conventional petroleum.

The most distinctive feature of the shale oils, nevertheless, is their high nitrogen content, which makes them unsuitable as a conventional feedstock for the refinery. Petroleum generally contains around 0.2% by weight of nitrogen, whereas crude shale oils contain generally of the order of 1% to approximately 3% by weight or more of nitrogen. Moreover, the nitrogen compounds present in petroleum are generally concentrated in relatively high boiling ranges, whereas the nitrogen of the compounds present in crude shale oils is generally distributed throughout all of the boiling ranges of the material. The nitrogen compounds in petroleum are primarily non-basic compounds, whereas, generally, around half of the nitrogen compounds present in crude shale oils are basic. These basic nitrogen compounds are particularly undesirable in refinery feedstocks, since these compounds often act as catalyst poisons. Furthermore, the stability of the products is a problem which is common to numerous products derived from shale oil. Such instability, including photosensitivity, appears to result essentially from the presence of nitrogen compounds. Consequently, crude shale oils must generally be subjected to a stringent refining treatment (high total pressure) in order to obtain a synthetic crude petroleum or fuel base products which meet the specifications in force.

It is also known that shale oils may contain numerous metal compounds in traces, generally present in the form of organometallic complexes. The metal compounds include the conventional contaminants such as nickel, vanadium, calcium, sodium, lead or iron, but also metal compounds of arsenic. Indeed, shale oils may contain an amount of arsenic of more than 20 ppm, whereas the amount of arsenic in crude petroleum is generally in the ppb (parts per billion) range. All of these metal compounds are catalyst poisons. More particularly, they irreversibly poison the hydrotreating catalysts and hydrogenating catalysts by gradually being deposited on the active surface. The conventional metal compounds and part of the arsenic are found primarily in heavy cuts, and are removed by deposition on the catalyst. On the other hand, when the products containing arsenic are capable of generating volatile compounds, these compounds may be found partly in the lighter cuts and may, as a result, poison the catalysts in subsequent converting processes, during refining or in petrochemistry.

Furthermore, shale oils generally contain sandy sediments originating from bituminous shale fields from which the shale oils are extracted. These sandy sediments may give rise to clogging problems, especially in fixed bed reactors.

Lastly, shale oils may contain waxes, which give them a pour point higher than the ambient temperature, thereby preventing their transport in oil pipelines.

In view of appreciable resources, and in view of their evaluation as being a promising source of petroleum, there exists a genuine need for converting shale oils into lighter products which can be utilized as fuels and/or raw materials for petrochemistry. Methods for converting shale oils are known. Conventionally, conversion is practised alternatively by coking, by hydrovisbreaking (thermal cracking in the presence of hydrogen) or by hydroconverting (catalytic hydrogenation). Liquid/liquid extraction processes are also known.

For instance, patent document CA2605056 describes a method for converting heavy feedstocks having an initial boiling point of greater than 340° C., wherein the feedstock is subjected to a deasphalting step and then the raffinate obtained is hydroconverted in the presence of a dispersed catalyst. The raffinate which is not hydroconverted is recycled to the deasphalting step. The feedstocks contemplated in this document are solely of the atmospheric residue and vacuum residue type obtained from conventional petroleum. A single desulphurizing step is envisaged on the atmospheric gas-oil cut. A further step of vacuum distillation with recycling of the residue under vacuum is shown.

Patent document CA2464796 describes a method for converting heavy feedstocks that is fairly similar to aforementioned CA2605056. In CA2464796, there is no reference to the treatment of shale oil; the feedstocks contemplated are just atmospheric residues and vacuum residues. Moreover, the vacuum distillation step is optional. Lastly, there is no recycling of the residue.

Patent application US2007138058 describes a slurry hydroconverting method in which the feedstock undergoes a preliminary step of hydrotreating or deasphalting. The converted products are subjected to an unspecified hydrofining operation. The process described is applied to various heavy feedstocks obtained from conventional petroleum. There is no mention of shale oils.

Patent application US2004163996 relates to a process for treating atmospheric residue or vacuum residue, in which said residue undergoes a deasphalting step and then the cuts of deasphalted oil (raffinate) and asphalt (extract) are separately hydroconverted by an ebullating bed process.

Patent application US 2009166253 envisages the linking of a number of deasphalting steps, followed by a step of hydrocracking of one or more deasphalted oils, after which they are fractionated. The process is applied to various feedstocks, including shale oils.

OBJECT OF THE INVENTION

The particular feature of shale oils in having a certain number of metallic and/or non-metallic impurities makes it much more complex to convert this non-conventional resource than petroleum. The challenge for the industrial development of methods for converting shale oils is therefore the need to develop methods which are suited to the feedstock, allowing the yield of high-quality fuel bases to be maximized. The conventional refining treatments known from petroleum must therefore be adapted to the specific composition of the shale oils.

The present invention aims to improve the known methods for converting hydrocarbon feedstocks comprising a shale oil by increasing, especially, the yield of fuel bases for a combination of steps having a specific linkage, and a treatment appropriate to each fraction obtained from the shale oils. Likewise, an object of the present invention is to obtain high-quality products having more particularly a low sulphur, nitrogen and arsenic content, preferably meeting the specifications. Another objective is to provide a method which is simple, i.e. having as few steps as necessary, while remaining effective, allowing capital investment costs to be limited.

In its broadest form, and according to a first aspect, the present invention is defined as a method for converting hydrocarbon feedstock comprising at least one shale oil having a nitrogen content of at least 0.1%, often at least 1% and very often at least 2% by weight, characterized in that it comprises the following steps:

a) The feedstock is subjected to a decontamination, to give a residue and a decontaminated oil,

b) The decontaminated oil is conveyed to a section for hydroconverting in the presence of hydrogen, said section comprising at least one ebullating bed reactor operating in gas and liquid upflow mode and containing at least one supported hydroconverting catalyst,

c) The effluent obtained in step b) is conveyed at least partly, and often entirely, into a fractionating zone, from which, by atmospheric distillation, a gaseous fraction, a naphtha fraction, a gas-oil fraction and a fraction heavier than gas-oil are recovered,

d) Said naphtha fraction is treated at least partly, and often entirely, in a first section for hydrotreating in the presence of hydrogen, said first section comprising at least one first fixed bed reactor containing at least one first hydrotreating catalyst, and

e) Said gas-oil fraction is treated at least partly, and often entirely, in a second section for hydrotreating in the presence of hydrogen, said second section comprising at least one second fixed bed reactor containing at least one second hydrotreating catalyst.

The hydroconverting section in step b) typically comprises from one to three, and preferably two, reactors in series, and the first and second hydrotreating sections in steps d) and e) typically comprise, each independently, from one to three reactors in series.

The method advantageously comprises a further step f) in which the fraction heavier than gas-oil is recycled to step a) in order to be decontaminated.

The research work carried out by the applicants into the conversion of shale oils has led to the finding that an improvement to the existing methods, in terms of yield of fuel bases and in terms of product purity, is possible through a combination of various steps linked in a specific way. Each fraction obtained by the method according to the invention is subsequently conveyed into a treatment section.

In a first phase, the hydrocarbon feedstock comprising shale oil is subjected to a decontamination, which advantageously takes the form of a deasphalting, by means of a solvent in the form of a C₃-C₅ hydrocarbon, alone or in a mixture. The solvent is selected for separating the asphalt, which contains heavy aromatic compounds which will not be soluble in the solvent. The extraction step also enables a reduction in the proportion of aromatic nitrogen compounds which are resistant to hydrodenitrogenation (denitrogenation by catalytic hydrogenation), the removal of the particles and minerals originating from the bituminous shale, and the entrainment of a large part of the heavy organometallic compounds containing metals such as vanadium or nickel. These heavy organometallic compounds may comprise porphyrin moieties within them, which are responsible for the sequestration of the metals. Deasphalting is carried out on the entirety of the feedstock, in order to maximize extraction of the aromatic nitrogen compounds present in cuts ranging from LPG to the fraction heavier than gas-oil. Using a deasphalting on the entirety of the feedstock prior to hydroconverting enables removal of a large part of the metallic poisons present in the feedstock, more particularly the arsenic, and also of the heavy nitrogen compounds, which are generally aromatic. Accordingly, the rate of renewal of the hydroconverting catalyst is reduced, and the hydrogen requirements are limited, owing to the reduced presence of compounds to be converted in the deasphalted oil. Examples of aromatic nitrogen compounds that may be encountered in feedstocks capable of being treated by the method of the invention include pyridines, pyrimidines, pyrazines, quinolines, isoquinolines, pyrroles, imidazoles, and other monocyclic or polycyclic aromatic nitrogen heterocycles. These aromatic nitrogen compounds may be substituted by various hydrocarbon chains, possibly containing heteroatoms such as nitrogen, oxygen and sulphur.

The second step comprises an ebullating bed hydroconversion. The technology of the ebullating bed, relative to the technology of the fixed bed, enables the treatment of feedstocks which are heavily contaminated with metals, heteroatoms and sediments, such as the shale oils, while exhibiting conversion rates which are generally greater than 50%. Indeed, in this second step, the shale oil is converted into molecules which enable the generation of future fuel bases. The majority of the metallic compounds, of the sediments and of the heterocyclic compounds is removed. The effluent emerging from the ebullating bed therefore contains the most resistant nitrogen and sulphur compounds, and possibly volatile arsenic compounds which are present in lighter components.

The effluent obtained in the hydroconverting step is subsequently fractionated by atmospheric distillation, producing various fractions, for which a treatment specific to each fraction is carried out subsequently. The atmospheric distillation enables the preparation, in a single step, of the various fractions desired (naphtha, gas-oil), thereby facilitating downstream hydrotreating adapted to each fraction and, consequently, the direct production of gas-oil or naphtha fuel base products which meet the various specifications. Fractionation after hydrotreating is therefore not necessary.

Owing to the high level of reduction in contaminants in the ebullating bed, the light fractions (naphtha and gas-oil) contain fewer contaminants and can therefore be treated in a fixed bed section, which generally has improved hydrogenation kinetics in relation to the ebullating bed. Similarly, the operating conditions can be milder because of the limited contaminants content. Providing a treatment for each fraction permits better operability in accordance with the desired products. Depending on the operating conditions selected (more or less stringent), it is possible to obtain either a fraction which can be conveyed to a fuels pool or a finished product which meets the specifications (sulphur content, smoke point, cetane, aromatics content, etc.) in force.

Upstream of the catalytic hydrotreating beds, the fixed bed hydrotreating sections preferably comprise specific guard beds for any arsenic compounds and silicon compounds contained within the diesel and/or naphtha fractions. The arsenic compounds, which have escaped the ebullating bed (because they are generally relatively volatile), are trapped in the guard beds, thus preventing poisoning of the downstream catalysts, and enabling production of highly arsenic-depleted fuel bases.

The atmospheric distillation also enables the concentration of the most resistant nitrogen compounds in the fraction which is heavier than the gas-oil fraction, and which, in step f), is conveyed to step a) in order to be deasphalted. Accordingly, this recycling step enables the suppression of the fraction heavier than gas-oil, and therefore enables the problems of utilization and of economic outlets for this fraction to be minimized. The fraction which is heavier than the gas-oil may be conveyed to step a) at least partly, preferably totally.

The asphalts find applications in the production of heavy fuel oils and bitumens.

DETAILED DESCRIPTION

Hydrocarbon Feedstock

The hydrocarbon feedstock comprises at least one shale oil or a mixture of shale oils. The term “shale oil” is used here in its broadest sense and is intended to include any shale oil or a shale oil fraction which contains nitrogenous impurities. This includes crude shale oil, whether obtained by pyrolysis, by solvent extraction or by other means, or shale oil which has been filtered to remove the solids, or which has been treated by one or more solvents, chemical products, or other treatments, and which contains nitrogenous impurities. The term “shale oil” also comprises the shale oil fractions obtained by distillation or by another fractionating technique.

The shale oils used in the present invention generally have a Conradson carbon content of at least 0.1% by weight and generally at least 5% by weight, an asphaltenes content (IP143 standard/with C7) of at least 1%, often at least 2% by weight. Their sulphur content is generally at least 0.1%, often at least 1% and very often at least 2%, and even up to 4% or even 7% by weight. The amount of metals they contain is generally at least 5 ppm by weight, often at least 50 ppm by weight, and typically at least 100 ppm by weight or at least 200 ppm by weight. Their nitrogen content is generally at least 0.5%, often at least 1% and very often at least 2% by weight. Their arsenic content is generally greater than 1 ppm by weight, and up to 50 ppm by weight.

The method according to the present invention is intended for converting shale oils. Nevertheless, the feedstock may further comprise, in addition to the shale oil, other, synthetic liquid hydrocarbons, more particularly those which contain a substantial amount of cyclic organic nitrogen compounds. This includes oils derived from coal, oils obtained on the basis of heavy tars, bituminous sands, pyrolysis oils from ligneous residues such as wood residues, crudes obtained from biomass (“biocrudes”), vegetable oils and animal fats.

Other hydrocarbon feedstocks may also supplement the shale oil. The feedstocks are selected from the group consisting of vacuum distillates and direct distillation residues, vacuum distillates and unconverted residues obtained from conversion processes, such as, for example, those originating from distillation to the point of coke (coking), products obtained from fixed-bed hydroconversion of heavy fractions, products obtained from ebullating-bed processes for hydroconversion of heavy fractions, and oils deasphalted using solvents (for example, oils deasphalted with propane, with butane and with pentane, originating from the deasphalting of vacuum residues from direct distillation or of vacuum residues obtained from hydroconversion processes). The feedstocks may further comprise light cycle oil (LCO) of various origins, heavy cycle oil (HCO) of various origins, and also gas-oil cuts which originate from catalytic cracking and have in general a distillation range from about 150° C. to about 650° C. The feedstocks may also comprise aromatic extracts obtained in the manufacture of lubricating oils. The feedstocks may also be prepared and used in a mixture, in any proportions.

Hydrocarbons added to shale oil or to the mixture of shale oils may represent from 20% to 60% by weight of the total feedstock (shale oil or mixture of shale oils+added hydrocarbons), or from 10% to 90% by weight.

Decontamination

The hydrocarbon feedstock comprising the shale oil is sent to a decontaminating step [step a)]. The objective in this step is to extract the contaminant-rich compounds. These contaminants contain heavy compounds rich in heteroatoms such as nitrogen, oxygen and sulphur, and also the major part of the metals present in the feedstock.

One preferred decontaminating method is deasphalting.

Such decontamination may be performed by extraction with a solvent or by ultrafiltration on a membrane, for example Zr on carbon, for example in the presence of supercritical CO₂ (C. Rodriguez et al. Desalination 144 (2002), 173-178).

Deasphalting by means of a solvent is carried out under conditions which are well known to the skilled person. Processes which may be used include, for example, the Solvahl process from Axens, or the Rose process from KBR.

Deasphalting traditionally takes place by means of an aliphatic solvent, generally a C₃-C₇, preferably C₃-C₅, alkane or cycloalkane, or mixtures thereof in any proportions.

The decontamination in step a) is carried out with a solvent selected from the group consisting of propane, n-butane, isobutane, n-pentane, cyclopentane, 2-methylbutane, 2,2-dimethylpropane, and mixtures thereof in any proportions.

Deasphalting may be performed by any means known to the skilled person. The extraction is generally performed in a mixer-settler or in an extraction column. The extraction is preferably performed in an extraction column.

The operating conditions are in general a solvent/feedstock ratio of 3/1 to 8/1 vol. %/vol. %, preferably of 4/1 to 6/1 vol. %/vol. %, and a temperature profile of between 60 and 250° C., preferably between 60° C. and 200° C. The operating pressure must be maintained at a value greater than the critical pressure of the solvent used. It is preferably between 4 and 5 MPa.

In one preferred embodiment, a mixture comprising the hydrocarbon feedstock and a first fraction of a solvent charge is introduced into the extraction column, the volume ratio between the solvent charge fraction and the hydrocarbon feedstock being called the level of solvent injected with the feedstock. The object in this step is to bring about effective mixing of the feedstock with the solvent entering the extraction column. In the settling zone at the bottom of the extractor, a second solvent charge fraction can be introduced, the volume ratio between the second charge fraction and the hydrocarbon feedstock being called the level of solvent injected at the bottom of the extractor. The volume of the hydrocarbon feedstock in question in the settling zone is generally that introduced into the extraction column. The sum of the two volume ratios between each of the solvent charge fractions and the hydrocarbon feedstock is called the overall level of solvent. Settling of the asphalt involves counter-current washing of the asphalt emulsion in the solvent+oil mixture with pure solvent. It is promoted by an increase in the level of solvent (the procedure in fact involves replacing the solvent+oil environment with a pure solvent environment) and by a reduction in the temperature.

The solvent can be recovered either by evaporation of the solvent (multiple flash stages with pressure lowering and temperature elevation) followed generally by steam stripping, or by operating the solvent separation procedure at a pressure greater than the critical pressure of the solvent or greater than that of the deasphalted oil/solvent mixture, thereby reducing the energy cost of the evaporation.

Extraction produces a feedstock of reduced nitrogen content, this content being approximately half in relation to the nitrogen content of the feedstock. At least part, and preferably the entirety, of the deasphalted oil is conveyed to an ebullating bed reactor for the purpose of hydroconversion in the presence of hydrogen. The ebullating bed comprises a supported hydroconversion catalyst.

According to one preferred variant, the asphalt is conveyed into an oxyvapogasification section in which it is converted to a gas containing hydrogen and carbon monoxide. This gas mixture can be used for the synthesis of methanol or for the synthesis of hydrocarbons by the Fischer-Tropsch reaction. This mixture, in the context of the present invention, is preferably conveyed into a “shift” conversion (steam conversion) section in which, in the presence of steam, it is converted into hydrogen and into carbon dioxide. The hydrogen obtained may be employed in steps b), d) and e) of the method according to the invention. The asphalt obtained in step a) may also be used as a solid fuel or, after fluxing, as a liquid fuel, or may form part of the composition of bitumens (after an optional blowing step) and/or of heavy fuel oils.

The deasphalting of the feedstock therefore enables extraction of resistant aromatic compounds containing nitrogen and of contaminants (metals), and also of the particles and sediments, which make up sometimes 0.2% by weight of the feedstock. The implementation of a preliminary step of deasphalting the feedstock enables retention of the hydroconversion catalyst, used in the following step, and minimization of the continuous top-ups of catalyst.

Hydroconversion

Following decontamination by deasphalting, the raffinate (deasphalted oil or DAO, or decontaminated feedstock) obtained is subjected to an ebullating-bed hydroconverting step [step b)]. By hydroconverting is meant reactions of hydrogenation, hydrotreating, hydrodesulphurization, hydrodenitrogenation, hydrodemetallation and hydrocracking.

The operation of the ebullating-bed catalytic reactor, including the recycling of the liquids from the reactor to the top through the agitated catalyst bed, is generally well known. Ebullating bed technologies use supported catalysts, generally in the form of extrudates having a diameter of generally of the order of 1 mm or less than 1 mm, for example greater than or equal to 0.7 mm. The catalysts remain inside the reactors and are not evacuated with the products. The catalytic activity can be held constant by virtue of on-line replacement (addition and withdrawal) of the catalyst. There is therefore no need to shut down the unit in order to change the spent catalyst, or to increase the reaction temperatures along the cycle in order to compensate for deactivation. Moreover, working with constant operating conditions enables consistent product qualities and consistent yields to be obtained throughout the cycle of the catalyst. Since the catalyst is held in agitation by substantial recycling of liquid, the head loss over the reactor remains low and constant, and the heat of reaction is rapidly averaged over the catalyst bed, which is therefore almost isothermal and does not require cooling via the injection of quenches. Implementing the hydroconversion in an ebullating bed obviates the problems of catalyst contamination that are associated with the deposits of impurities that are present naturally in shale oils.

The conditions in step b) of treating the feedstock in the presence of hydrogen are customarily conventional conditions for ebullating-bed hydroconversion of a liquid hydrocarbon fraction. It is customary to operate under a total pressure of 2 to 35 MPa, preferably of 10 to 20 MPa, at a temperature of 300° C. to 550° C. and often of 400° C. to 450° C. The hourly space velocity (HSV) and the hydrogen partial pressure are important factors, which are selected according to the characteristics of the product to be treated and to the desired conversion. The HSV is usually situated within a range from 0.2 h⁻¹ to 1.5 h⁻¹ and preferably from 0.4 h⁻¹ to 1 h⁻¹. The amount of hydrogen mixed with the feedstock is customarily from 50 to 5000 normal cubic metres (Nm³) per cubic metre (m³) of liquid feedstock, and usually from 100 to 1000 Nm³/m³, and preferably from 300 to 500 Nm³/m³.

This hydroconverting step b) may usually be implemented under the conditions of the T-STAR® process, as described for example in the article Heavy Oil Hydroprocessing, published by the AlChE, Mar. 19-23, 1995, Houston, Tex., paper number 42d. It may also be implemented under the conditions of the H-OIL® process, as described for example in the article published by NPRA, Annual Meeting, Mar. 16-18, 1997, J. J. Colyar and L. I. Wisdom under the title The H-Oil®Process, A Worldwide Leader In Vacuum Residue Hydroprocessing.

The hydrogen required for the hydroconversion (and for the subsequent hydrotreating operations) may come from the steam reforming of hydrocarbons (methane) or else from the gas obtained from bituminous shales during the production of shale oils.

The catalyst in step b) is preferably a conventional granular hydroconversion catalyst, comprising, on an amorphous support, at least one metal or metal compound having a hydrodehydrogenating function. Generally speaking, a catalyst is used whose pore distribution is suitable for the treatment of feedstocks containing metals.

The hydrodehydrogenating function may be provided by at least one group VIII metal selected from the group consisting of nickel and/or cobalt, optionally in combination with at least one group VIB metal selected from the group consisting of molybdenum and/or tungsten. It is possible, for example, to use a catalyst comprising from 0.5% to 10% by weight of nickel and preferably from 1% to 5% by weight of nickel (expressed as nickel oxide, NiO) and from 1% to 30% by weight of molybdenum, preferably from 5% to 20% by weight of molybdenum (expressed as molybdenum oxide, MoO₃), on an amorphous inorganic support. The total amount of oxides of metals from groups VIB and VIII is often from 5% to 40% by weight and generally from 7% to 30% by weight. The weight ratio expressed as metal oxide between group VI metal (or metals) and group VIII metal (or metals) is generally from 20 to 1 and usually from 10 to 2.

The support of the catalyst will be selected, for example, from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also include other compounds, for example oxides selected from the group consisting of boron oxide, zirconia, titanium oxide and phosphoric anhydride. It is usual to use an alumina support, and very often an alumina support doped with phosphorus and optionally with boron. In this case, the concentration of phosphoric anhydride, P₂O₅, is customarily less than about 20% by weight and usually less than about 10% by weight, and at least 0.001% by weight. The concentration of boron trioxide, B₂O₃, is customarily from approximately 0% to approximately 10% by weight. The alumina used is customarily a γ (gamma) or η (eta) alumina. This catalyst is usually in the form of an extrudate. The catalyst in step b) is preferably based on nickel and molybdenum, doped with phosphorus and supported on alumina. Use may be made, for example, of an HTS 458 catalyst sold by Axens.

Prior to the injection of the deasphalted oil, the catalysts used in the method according to the present invention may undergo a sulphurizing treatment to convert at least partly the metallic species into sulphides before they are contacted with the feedstock to be treated. This activation treatment by sulphurization is well known to the skilled person and may be carried out by any method already described in the literature, either in situ, i.e. within the reactor, or ex situ.

The spent catalyst is partly replaced with fresh catalyst by withdrawal at the bottom of the reactor and introduction at the top of the reactor of fresh or new catalyst at regular intervals, for example by individual or quasi-continuous addition. It is possible, for example, to introduce fresh catalyst every day. The level of replacement of the spent catalyst by fresh catalyst may be, for example, from approximately 0.05 kg to approximately 10 kg per m³ of feedstock. This withdrawal and this replacement are carried out using devices which allow the continuous operation of this hydroconverting step. The unit customarily comprises a recirculation pump for maintaining the catalyst in an ebullating bed by continuous recycling of at least part of the liquid withdrawn at the top of the reactor and reinjected at the bottom of the reactor. It is also possible to convey the spent catalyst withdrawn from the reactor into a regenerating zone, in which the carbon and sulphur it contains are removed, and then to convey this regenerated catalyst back into the hydroconversion reactor in step b).

The operating conditions coupled with the catalytic activity allow feedstock conversion rates of possibly from 50% to 95%, preferably from 70% to 95%, to be obtained. The aforementioned degree of conversion is defined as the mass fraction of the feedstock at the start of the reaction section minus the mass fraction of the heavy fraction having a boiling point of more than 343° C. at the end of the reaction section, this figure being divided by the mass fraction of the feedstock at the start of the reaction section.

The technology of the ebullating bed allows treatment of feedstocks which are highly contaminated with metals, sediments and heteroatoms, without facing head loss problems or clogging problems, which are known when a fixed bed is used. The metals, such as nickel, vanadium, iron and arsenic, are largely removed from the feedstock by deposition on the catalysts during the reaction. The remaining (volatile) arsenic will be removed in the hydrotreating steps by specific guard beds. The sediments present in the shale oils are also removed via the replacement of the catalyst in the ebullating bed without disrupting the hydroconversion reactions. These steps also enable the removal, by hydrodenitrogenation, of the major part of the nitrogen, leaving only the most resistant nitrogen compounds.

The hydroconversion in step b) enables an effluent to be obtained that contains not more than 3000 ppm, preferably not more than 2000 ppm, by weight of nitrogen.

Fractionation by Atmospheric Distillation

The effluent obtained in hydroconverting step b) is conveyed at least partly, and preferably in its entirety, into a fractionating zone, from which a gaseous fraction, a naphtha fraction, a gas-oil fraction and a fraction heavier than the gas-oil fraction are recovered by atmospheric distillation.

The effluent obtained in step b) is preferably fractionated by atmospheric distillation into a gaseous fraction having a boiling point of less than 50° C., a naphtha fraction boiling at between about 50° C. and 150° C., a gas-oil fraction boiling at between about 150° C. and 370° C., and a fraction which is heavier than the gas-oil fraction and which boils generally at above 340° C., preferably at above 370° C.

The naphtha and diesel fractions are subsequently conveyed separately into hydrotreating sections. The fraction heavier than the gas-oil fraction is conveyed to the decontaminating unit of step a), where it is mixed into the feedstock.

The gaseous fraction contains gases (H₂, H₂S, NH₃, H₂O, CO₂, CO, C₁-C₄ hydrocarbons, etc.). It may advantageously undergo a purifying treatment for recovery of the hydrogen and its recycling into the hydroconverting section in step b) or into the hydrotreating sections in steps d) and e). Following purifying treatments, the C₃ and C₄ hydrocarbons may be used to form LPG (liquefied petroleum gas) products. The uncondensable gases (C₁-C₂) are generally used as internal fuel for the heating ovens of the hydroconversion and/or hydrotreating reactors. The C₁-C₄ hydrocarbons isolated from the gaseous fraction may optionally be used for carrying out decontaminating step a). The operating conditions of the decontamination (pressure, temperature, solvent/feedstock ratio) will vary with the nature of the feedstock containing a shale oil, and will be adjusted for compatibility with the desired purification parameters.

Hydrotreating of the Naphtha Fraction and of the Gas-Oil Fraction

The naphtha and gas-oil fractions are subsequently subjected separately to fixed-bed hydrotreating [steps d) and e)]. Hydrotreating refers to reactions of hydrodesulphurization, hydrodenitrogenation and hydrodemetallation. The objective, depending on the operating conditions, which are selected so as to be more or less stringent, is to bring the various cuts up to the specifications (sulphur content, smoke point, cetane, aromatics content, etc.) or to produce a synthetic crude petroleum. Treating the naphtha fraction in one hydrotreating section and the gas-oil fraction in another hydrotreating section allows improved operability in terms of the operating conditions, so as to be able to bring each cut up to the required specifications with a maximum yield and in a single step per cut. In this way, fractionation after hydrotreating is unnecessary. The difference between the two hydrotreating sections is based more on differences in operating conditions than on the selection of the catalyst.

The fixed-bed hydrotreating sections preferably comprise, upstream of the catalytic hydrotreating beds, specific guard beds for the arsenic compounds (arsenic-containing compounds) and silicon compounds that are optionally present in the naphtha and/or diesel fractions. The arsenic-containing compounds which have escaped the ebullating bed (being generally relatively volatile) are trapped in the guard beds, thereby preventing the poisoning of downstream catalysts and enabling highly arsenic-depleted fuel bases to be obtained.

The guard beds which allow removal of arsenic and silicon from naphtha or gas-oil cuts are known to the skilled person. They comprise, for example, an absorbent material comprising nickel deposited on an appropriate support (silica, magnesia or alumina) as described in FR2617497, or else an absorbent material comprising copper on a support, as described in FR2762004. Mention may also be made of the guard beds sold by Axens: ACT 979, ACT 989, ACT 961, ACT 981.

The operating conditions in each hydrotreating section are adapted to the feedstock to be treated. The operating conditions for hydrotreating the naphtha fraction are generally gentler than those for the gas-oil fraction.

In the naphtha fraction hydrotreating step [step d)] it is customary to operate under an absolute pressure of 4 to 15 MPa, often of 10 to 13 MPa. The temperature during this step d) is customarily from 280° C. to 380° C., often from 300° C. to 350° C. This temperature is customarily adjusted in accordance with the desired level of hydrodesulphurization. The hourly space velocity (HSV) is usually situated within a range from 0.1 h⁻¹ to 5 h⁻¹, and preferably from 0.5 h⁻¹ to 1 h⁻¹. The amount of hydrogen mixed with the feedstock is customarily from 100 to 5000 normal cubic metres (Nm³) per cubic metre (m³) of liquid feedstock, and usually from 200 to 1000 Nm³/m³, and preferably from 300 to 500 Nm³/m³. It is useful to operate in the presence of hydrogen sulphide (for the sulphurizing of the catalyst), and the hydrogen sulphide partial pressure is customarily from 0.002 times to 0.1 times, and preferably from 0.005 times to 0.05 times, the total pressure.

In the gas-oil fraction hydrotreating step [step e)] it is customary to operate under an absolute pressure of 7 to 20 MPa, often of 10 to 15 MPa. The temperature during this step e) is customarily from 320° C. to 450° C., often from 340° C. to 400° C. This temperature is customarily adjusted depending on the desired level of hydrodesulphurization. The mass hourly velocity is between 0.1 and 1 h⁻¹. The hourly space velocity (HSV) is usually situated within a range from 0.2 h⁻¹ to 1 h⁻¹, and preferably from 0.3 h⁻¹ to 0.8 h⁻¹. The amount of hydrogen mixed into the feedstock is customarily from 100 to 5000 normal cubic metres (Nm³) per cubic metre (m³) of liquid feedstock, and usually from 200 to 1000 Nm³/m³, and preferably from 300 to 500 Nm³/m³. It is useful to operate in the presence of hydrogen sulphide, and the hydrogen sulphide partial pressure is customarily from 0.002 times to 0.1 times, and preferably from 0.005 times to 0.05 times, the total pressure.

In the hydrotreating sections, the ideal catalyst must have a high hydrogenating power, so as to produce thorough refining of the products, and to obtain a substantial lowering of the sulphur content and nitrogen content. In the preferred embodiment, the hydrotreating sections operate at relatively low temperature, which promotes thorough hydrogenation and a limitation on the coking of the catalyst. The use of a single catalyst or of two or more different catalysts, simultaneously or successively, in the hydrotreating sections would not depart from the scope of the present invention. The hydrotreating in steps d) and e) is customarily carried out industrially in one or more reactors with liquid downflow.

In the two hydrotreating sections [steps d) and e)], the same type of catalyst is used; the catalysts in each section may be identical or different. At least one fixed bed of conventional hydrotreating catalyst is used, comprising, on an amorphous support, at least one metal or metal compound having a hydrodehydrogenating function.

The hydrodehydrogenating function may be provided by at least one group VIII metal selected from the group consisting of nickel and/or cobalt, optionally in combination with at least one group VIB metal selected from the group consisting of molybdenum and/or tungsten. It is possible, for example, to use a catalyst comprising from 0.5% to 10% by weight of nickel and preferably from 1% to 5% by weight of nickel (expressed as nickel oxide, NiO) and from 1% to 30% by weight of molybdenum, preferably from 5% to 20% by weight of molybdenum (expressed as molybdenum oxide, MoO₃), on an amorphous inorganic support. The total amount of oxides of metals from groups VI and VIII is often from about 5% to about 40% by weight, and generally from about 7% to 30% by weight, and the weight ratio expressed in terms of metal oxide between metal (or metals) from group VIB to metal (or metals) from group VIII is in general from about 20 to about 1, and usually from about 10 to about 2.

The support is for example selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also include other compounds, for example oxides selected from the group consisting of boron oxide, zirconia, titanium oxide and phosphoric anhydride. It is usual to use an alumina support, and very often an alumina support doped with phosphorus and optionally with boron. In this case, the concentration of phosphoric anhydride, P₂O₅, is customarily less than about 20% by weight and usually less than about 10% by weight, and is at least 0.001% by weight. The concentration of boron trioxide, B₂O₃, is customarily from approximately 0% to approximately 10% by weight. The alumina used is customarily a γ (gamma) or η (eta) alumina. This catalyst is usually in the form of beads or extrudates.

Prior to the injection of the feedstock, the catalysts used in the method according to the present invention are preferably subjected to a sulphurizing treatment enabling to convert at least partly the metallic species into sulphides before they are contacted with the feedstock to be treated. This activation treatment by sulphurization is well known to the skilled person and may be carried out by any method already described in the literature, either in situ, i.e. within the reactor, or ex situ.

The hydrotreating in step d) of the naphtha cut produces a cut containing not more than 1 ppm by weight of nitrogen, preferably not more than 0.5 ppm of nitrogen, and not more than 5 ppm by weight of sulphur, preferably not more than 0.5 ppm of sulphur.

The hydrotreating in step e) of the gas-oil cut produces a cut containing not more than 100 ppm of nitrogen, preferably not more than 20 ppm of nitrogen, and not more than 50 ppm of sulphur, preferably not more than 10 ppm of sulphur.

Catalytic Cracking

Finally, according to one variant, the method according to the invention may comprise a catalytic cracking step [step g)], in which at least part, and preferably the entirety, of the fraction heavier than gas-oil, obtained in step c), is conveyed into a conventional catalytic cracking section, in which said fraction heavier than gas-oil is treated conventionally, under conditions well known to the skilled person, to produce a second gaseous fraction, a second petrol fraction, a second gas-oil fraction and a second fraction heavier than gas-oil, referred to as “slurry”. The second gas-oil fraction will for example be conveyed at least partly to fuel reservoirs (pools) and/or recycled, at least partly, or even in its entirety, to the gas-oil hydrotreating step e). The second fraction heavier than gas-oil will, for example, be at least partly, or even in its entirety, conveyed to the heavy fuel oil reservoir (pool) and/or recycled at least partly, or even in its entirety, to the catalytic cracking step g). At least part of the second fraction heavier than gas-oil, obtained at the end of step g), is recycled to the decontaminating step a). In the context of the present invention, the expression “conventional catalytic cracking” encompasses cracking processes which comprise at least one step of catalyst regeneration by partial combustion, and those which comprise at least one step of catalyst regeneration by total combustion, and/or those comprising both at least one partial combustion step and at least one total combustion step.

For example, a summary description of catalytic cracking (the first industrial implementation of which goes back to 1936 (Houdry process) or 1942 for the use of fluidized bed catalyst) will be found in Ullmans Encyclopedia of Industrial Chemistry Volume A 18, 1991, pages 61 to 64. It is customary to use a conventional catalyst comprising a matrix, optionally an additive and at least one zeolite. The amount of zeolite is variable but is customarily from about 3% to 60% by weight, often from about 6% to 50% by weight and usually from about 10% to 45% by weight. The zeolite is customarily dispersed in the matrix. The amount of additive is customarily from about 0% to about 30% by weight. The amount of matrix represents the rest up to 100% by weight. The additive is generally selected from the group consisting of the oxides of metals from group IIA of the periodic table of the elements, such as, for example, magnesium oxide or calcium oxide, the oxides of rare earths, and the titanates of the metals from group IIA. The matrix is usually a silica, an alumina, a silica-alumina, a silica-magnesia, a clay or a mixture of two or more of these products. The zeolite most commonly used is zeolite Y. Cracking is carried out in a substantially vertical reactor in either upflow or downflow mode. The selection of the catalyst and of the operating conditions are dependent on the target products in dependence on the feedstock treated, as is described, for example, in the article by M. Marcilly, pages 990-991, published in the journal of the Institut Français du Pétrole, November-December 1975, pages 969-1006. It is customary to operate at a temperature from 450° C. to 600° C. and with reactor residence times of less than 1 minute, often from about 0.1 to about 50 seconds.

The catalytic cracking step g) may also be a fluidized bed catalytic cracking step, for example according to the process called R2R. This step may be performed conventionally as known to skilled persons under appropriate cracking conditions for producing hydrocarbon products with a lower molecular weight. Descriptions of operation and of catalysts which can be used in the context of fluidized bed cracking in this step g) are described for example in the patent documents U.S. Pat. No. 5,286,690, U.S. Pat. No. 5,324,696 and EP-A-699224.

The fluidized bed catalytic cracking reactor may operate in upflow mode or in downflow mode. Although not a preferred embodiment of the present invention, it is likewise possible to contemplate performing the catalytic cracking in a moving bed reactor. Particularly preferred catalytic cracking catalysts are those containing at least one zeolite, customarily in a mixture with an appropriate matrix such as, for example, alumina, silica or silica-alumina.

According to a penultimate aspect, the invention relates to the preparation of a synthetic crude by a method according to one of its preceding aspects.

According to a last aspect, the invention relates to a plant intended for treating a shale oil, employing a method according to one of its preceding aspects.

Such a Plant Comprises:

-   -   a section for decontaminating the shale oil to be treated,     -   a section for hydroconverting in the presence of hydrogen,         comprising an ebullating bed reactor operating in gas and liquid         upflow mode and containing at least one supported         hydroconverting catalyst,     -   a zone for fractionation by atmospheric distillation,     -   a section for hydrotreating in the presence of hydrogen,         comprising a fixed bed reactor containing at least one         hydrotreating catalyst,     -   another section for hydrotreating in the presence of hydrogen,         comprising at least one fixed bed reactor containing at least         one hydrotreating catalyst,

These elements are arranged for the implementation of the method according to the invention.

Accordingly, for example:

-   -   the section for decontaminating is connected to the section for         hydroconverting in order to feed this section for         hydroconverting with treated shale oil issued from the section         for decontaminating,     -   the section for hydroconverting is connected to the zone for         fractionation in order to feed this zone for fractionation with         effluents issued from the section for hydroconverting,     -   one duct (line) connects the zone for fractionation to one of         the two sections for hydrotreating, another duct (line)         connecting the zone for fractionation to the other of the two         sections for hydrotreating.

A section for catalytic cracking may be provided, another duct (line) connecting the zone for fractionation to this section for catalytic cracking.

The plant may further comprise one or several recycle ducts for conveying different fractions to the section for decontaminating or to the section for catalytic cracking.

FIG. 1 represents diagrammatically the method according to the present invention.

FIG. 2 represents diagrammatically a variant of the method which includes the catalytic cracking step.

According to FIG. 1, the feedstock comprising the shale oil (1) to be treated enters by the line (2) into a decontaminating section (3). This step (3) is a deasphalting carried out using a solvent (not shown), which produces a deasphalted oil (4) and a residue (5). The residue (5), via the line (6), may be used as fuel or may feed a gasification unit for producing hydrogen and energy.

The deasphalted oil (4) is conveyed by a line (7) to an ebullating bed hydroconverting section (8), in the presence of hydrogen (9), the hydrogen (9) being introduced by the line (10). The effluent from the ebullating bed hydroconverting section (8) is conveyed by the line (11) into an atmospheric distillation column (12), at the end of which a gaseous fraction (13), a naphtha fraction (14), a gas-oil fraction (15) and a fraction (16) heavier than the gas-oil fraction are recovered. The gaseous fraction (13), containing hydrogen, may be purified (not shown) for recycling the hydrogen and reinjecting it into the ebullating bed hydroconverting section (8) via the line (10), and/or into the hydrotreating sections (17) and/or (18) via the lines (19) and (20). The naphtha fraction (14) is conveyed into the fixed bed hydrotreating section (17), at the end of which a naphtha fraction (21) depleted in impurities is recovered. The gas-oil fraction (15) is conveyed into the fixed bed hydrotreating section (18), at the end of which a gas-oil fraction (22) depleted in impurities is recovered. The two hydrotreating sections (17) and (18) are fed by hydrogen via the lines (23) and (24). The fraction (16) heavier than the gas-oil fraction is recycled within the deasphalting section (3) via the line (25).

In FIG. 2, the liquid/liquid extracting, hydroconverting, separating and hydrotreating steps (and reference symbols) are identical to those of FIG. 1. The fraction (16) heavier than the gas-oil fraction, emerging from the atmospheric distillation step, may be sent to a catalytic cracking section (26) by means of a line (27). The effluent from this section (26) is sent via the line (28) to a fractionating section (29), preferably an atmospheric distillation, from which a fuels or middle distillates fraction is recovered, comprising at least one second petrol fraction (30), a second gas-oil fraction (31) and a second fraction (32) heavier than gas-oil, also called “slurry”. The second gas-oil fraction (31) is conveyed at least partly to the fuel reservoirs (pools) and/or is recycled at least partly, or even in its entirety, to the gas-oil hydrotreating step e) (18) via the line (33). The second fraction (32) heavier than gas-oil (“slurry”) is for example, at least partly or even in its entirety, conveyed to the heavy fuel-oil reservoir (pool) and/or is recycled, at least partly or even in its entirety, to the catalytic cracking step (26) via the line (34). The second fraction (32) heavier than gas-oil is optionally conveyed partly or in its entirety to the deasphalting unit (3) by a line (35).

EXAMPLE

A shale oil is treated that has the characteristics set out in Table 1.

TABLE 1 Characteristics of the shale oil feedstock Density 15/4 — 0.948 Hydrogen % by weight 10.9 Sulphur % by weight 1.9 Nitrogen % by weight 0.9 Oxygen % by weight 2.4 Asphaltenes % by weight 2.0 Conradson carbon % by weight 3.1 Metals ppm 195

The shale oil, from which the hydrocarbons lower than C₅ have been separated, is subjected to a preliminary atmospheric distillation, from which volatile products, constituents of gas-oils, kerosenes and naphthas, are isolated and treated in the refining chain according to their properties. The atmospheric distillation of the shale oil also produces an atmospheric residue, which comprises within it a majority of compounds having a boiling point of greater than 350° C. The atmospheric residue, in the context of the present example, represents 40% by weight of the feedstock. This feedstock is subjected to propane deasphalting with a solvent/feedstock ratio of 5/1, at a temperature of 100° C. and at 4 MPa. This produces a deasphalted oil and a residue.

The deasphalted oil is treated in an ebullating bed reactor containing the commercial catalyst HTS458 from Axens. The operating conditions are as follows:

-   -   Temperature in the reactor: 425° C.     -   Pressure: 195 bar (19.5 MPa)     -   Hydrogen/feedstock ratio: 400 Nm³/m³     -   Overall HSV: 0.6 h⁻¹

The liquid products obtained from the reactor are fractionated by atmospheric distillation to give a naphtha fraction (C5+−150° C.), a gas-oil fraction (150-370° C.) and a residual fraction 370° C.+, which constitutes a fraction heavier than gas-oil.

The naphtha fraction is subjected to fixed bed hydrotreating using an NiMo-on-alumina catalyst. The operating conditions are as follows:

-   -   Temperature in the reactor: 320° C.     -   Pressure: 50 bar (5 MPa)     -   Hydrogen/feedstock ratio: 400 Nm³/m³     -   Overall HSV: 1 h⁻¹

The gas-oil fraction is subjected to fixed bed hydrotreating using an NiMo-on-alumina catalyst. The operating conditions are as follows:

-   -   Temperature in the reactor: 350° C.     -   Pressure: 120 bar (12 MPa)     -   Hydrogen/feedstock ratio: 400 Nm³/m³     -   Overall HSV: 0.6 h⁻¹

The fraction heavier than gas-oil is subsequently subjected to catalytic cracking using a catalyst containing 20% by weight of zeolite Y and 80% by weight of a silica-alumina matrix. This feedstock, preheated to 135° C., is contacted at the bottom of a vertical reactor with a catalyst from a regenerator, the catalyst having been regenerated under hot conditions. The entry temperature of the catalyst into the reactor is 720° C. The ratio of the catalyst flow rate to the feedstock flow rate is 6.0. The calorific input of the catalyst at 720° C. enables the evaporation of the feedstock and the cracking reaction, both of which are endothermic. The average residence time of the catalyst in the reaction zone is approximately 3 seconds. The operating pressure is 1.8 bar absolute. The catalyst temperature measured at the end of the upwardly driven (riser) fluidized bed reactor is 525° C. The cracked hydrocarbons and the catalyst are separated by virtue of cyclones situated in a stripping zone (stripper) in which the catalyst is stripped. The catalyst, which has become loaded with coke during the reaction and then has been stripped in the stripping zone, is subsequently conveyed into the regenerator. The coke content of the solid (delta coke) at the start of the regenerator is 0.85%. This coke is burnt by air injected into the regenerator. The combustion, which is very exothermic, raises the temperature of the solid from 525° C. to 720° C. The hot regenerated catalyst emerges from the regenerator and is conveyed back to the bottom of the reactor.

The hydrocarbons separated from the catalyst emerge from the stripping zone. They are sent to a main fractionating tower, from which the gases and petrol cuts emerge at the top, and then, at the bottom of the tower, in order of increasing boiling point, the LCO and HCO cuts and the slurry (370° C.+) emerge.

Table 2 gives the properties of the various feedstocks in each step and also the yields obtained in the various units, and the overall yield. Hence it is observed that, starting from 100% by weight of shale oil, 86.2% by weight of products (LPG, naphtha, middle distillates) are obtained conforming to the commercial Euro V specifications.

TABLE 2 Refining unit Atmospheric Solvent Solvent H-Oil_(DC) FCC distillation deasphalting deasphalting Feedstock Shale oil Shale oil Shale oil Deasphalted H-Oil_(DC) C5+ 360° C.+ 360° C.+ oil bottoms Product Shale oil Deasphalted H-Oil_(DC) Products 360° C.+ oil Residue bottoms ex-FCC Yield of product % by weight 40.0 28.0 12.0 5.9 based on feedstock obtained from shale oil Properties of the products Density (d15/4) — 1.035 0.980 1.191 0.905 Sulphur % by weight 1.5 1.3 2.0 0.06 Total nitrogen % by weight 1.1 0.7 1.8 0.15 Yields over each unit LPG % by weight 2.0 10.0 Naphtha % by weight 19.0 55.0 Middle distillate % by weight 56.0 14.0 Bottoms % by weight 21.0 5.0 % by weight Yield over % by weight shale oil LPG % by weight 0.6 0.6 Naphtha % by weight 5.3 3.2 Middle distillate % by weight 15.7 0.8 Bottoms % by weight 5.9 0.3 Total 60.0 21.6 4.6 (LPG + naphtha + middle distillate)/shale oil feedstock Total 86.2 (LPG + naphtha + middle distillate)/shale oil feedstock 

1. Method for converting a shale oil or a mixture of shale oils having a nitrogen content of at least 0.1%, often at least 1% and very often at least 2% by weight, characterized in that it comprises the following steps: a) The feedstock is subjected to a decontamination, to give a residue and a decontaminated oil, b) The decontaminated oil is conveyed to a section for hydroconverting in the presence of hydrogen, said section comprising at least one ebullating bed reactor operating in gas and liquid upflow mode and containing at least one supported hydroconverting catalyst, c) The effluent obtained in step b) is conveyed at least partly, and often entirely, into a fractionating zone, from which, by atmospheric distillation, a gaseous fraction, a naphtha fraction, a gas-oil fraction and a fraction heavier than gas-oil are recovered, d) Said naphtha fraction is treated at least partly, and often entirely, in another section for hydrotreating in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreating catalyst, and e) Said gas-oil fraction is treated at least partly, and often entirely, in a section for hydrotreating in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreating catalyst.
 2. Method according to claim 1, characterized in that it further comprises a step f), in which at least a part of the fraction heavier than gas-oil is conveyed to step a) to be decontaminated.
 3. Method according to claim 1, wherein the effluent obtained in step b) is fractionated by atmospheric distillation into a gaseous fraction having a boiling point of less than 50° C., a naphtha fraction boiling at between about 50° C. and 150° C., a gas-oil fraction boiling at between about 150° C. and 370° C., and a fraction which is heavier than the gas-oil fraction and which boils generally at above 340° C., preferably above 370° C.
 4. Method according to claim 1, wherein the decontamination in step a) is carried out with a solvent selected from the group consisting of propane, n-butane, isobutane, n-pentane, cyclopentane, 2-methylbutane, 2,2-dimethylpropane, and mixtures thereof in any proportions.
 5. Method according to claim 1, wherein decontaminating step a) is carried out with a solvent/feedstock ratio of 3/1 to 8/1, preferably of 4/1 to 6/1, at a temperature of between 60° C. and 250° C., preferably between 60° C. and 200° C., and at a pressure of between 4 MPa and 5 MPa.
 6. Method according to claim 1, wherein the fixed bed hydrotreating section in step d) and/or e) comprises, upstream of the catalytic hydrotreating beds, at least one specific guard bed for arsenic compounds and silicon compounds.
 7. Method according to claim 1, wherein at least a part of the fraction heavier than gas-oil is conveyed into a catalytic cracking section, called step g), in which it is treated under conditions enabling production of a second gaseous fraction, a second petrol fraction, a second gas-oil fraction and a second fraction heavier than gas-oil.
 8. Method according to claim 7, wherein at least part of the second fraction heavier than gas-oil, obtained at the end of step g), is recycled to the start of said step g).
 9. Method according to claim 7, wherein at least part of the second gas-oil fraction, obtained at the end of step g), is recycled to gas-oil hydrotreating step e).
 10. Method according to claim 7, wherein at least part of the second fraction heavier than gas-oil, obtained at the end of step g), is recycled to decontaminating step a).
 11. Method according to claim 1, wherein hydroconverting step b) operates at a temperature of between 300° C. and 550° C., preferably between 400° C. and 450° C., at a total pressure of between 2 and 35 MPa, preferably of between 10 and 20 MPa, at a mass hourly velocity ((t of feedstock/h)/t of catalyst) of between 0.2 and 1.5 h⁻¹, and at a hydrogen/feedstock ratio of between 50 and 5000 Nm³/m³, preferably between 100 and 1000 Nm³/m³.
 12. Method according to claim 1, wherein step d) of hydrotreating the naphtha fraction operates at a temperature of between 280° C. and 380° C., preferably between 300° C. and 350° C., at a total pressure of between 4 and 15 MPa, preferably of between 10 and 13 MPa, at a mass hourly velocity of between 0.1 and 5 h⁻¹, preferably between 0.5 and 1 h⁻¹, and at a hydrogen/feedstock ratio of between 100 and 5000 Nm³/m³, preferably between 100 and 1000 Nm³/m³.
 13. Method according to claim 1, wherein step e) of hydrotreating the gas-oil fraction operates at a temperature of between 320° C. and 450° C., preferably between 340° C. and 400° C., at a total pressure of between 7 and 20 MPa, preferably of between 10 and 15 MPa, at a mass hourly velocity of between 0.1 and 1 h⁻¹, preferably between 0.3 and 0.8 h⁻¹, and at a hydrogen/feedstock ratio of between 100 and 5000 Nm³/m³, preferably between 200 and 1000 Nm³/m³.
 14. Method according to claim 1, wherein the catalysts in hydroconverting step b) and hydrotreating steps d) and e) are independently selected from the group of catalysts comprising a group VIII metal selected from the group consisting of Ni and/or Co, optionally a group VIB metal selected from the group consisting of Mo and/or W, on an amorphous support selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures thereof.
 15. Method according to claim 1, wherein the shale oil or the mixture of shale oils is supplemented by a hydrocarbon feedstock selected from the group consisting of oils derived from coal, oils obtained from heavy tars and bituminous sands, vacuum distillates, and residues of direct distillation, vacuum distillates and unconverted residues obtained from a residue conversion process, oils deasphalted with solvents, light cycle oils, heavy cycle oils, gas-oil cuts originating from catalytic cracking and having generally a distillation range from approximately 150° C. to approximately 650° C., aromatic extracts obtained in the manufacture of lubricating oils, pyrolysis oils of ligneous residues such as wood residues, crudes obtained from biomass (“biocrudes”), vegetable oils and animal fats, or mixtures of such feedstocks.
 16. Synthetic crude obtained by a method according to claim
 1. 17. Plant for treating a shale oil, comprising: a section for decontaminating the shale oil to be treated, a section for hydroconverting in the presence of hydrogen, comprising an ebullating bed reactor operating in gas and liquid upflow mode and containing at least one supported hydroconverting catalyst, a zone for fractionation by atmospheric distillation, a section for hydrotreating in the presence of hydrogen, comprising a fixed bed reactor containing at least one hydrotreating catalyst, another section for hydrotreating in the presence of hydrogen, comprising at least one fixed bed reactor containing at least one hydrotreating catalyst, these elements being arranged for the implementation of the method according to claim
 1. 